Performance indicators for ammonia based system analysis

https://doi.org/10.1016/j.seppur.2021.118959

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The performance assessment of the advanced configurations and the comparison among different PFD’s have been carried out with reference to the following process performance indicators:

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Energy consumption is considered as an indicator of the operating costs of the capture process. The specific equivalent work, $ω$ [ $MJ {kg}_{{CO}_{2} captured}^{- 1}$ ], considers the steam demand in the reboilers of the process as well as the electricity required for chilling, cooling, pumping and compression. The conversion of each energy consumer to specific equivalent work has been detailed in our previous work [14]. In addition, the specific equivalent work associated with the NH₃ capture loop defined by the FG-WW column and the NH₃ desorber, $ω_{FG−WW}^{{NH}_{3}}$ [ $MJ {kg}_{{NH}_{3} captured}^{- 1}$ ], has also been used as energetic performance indicator of those configurations aiming at controlling NH₃ emissions in the CO₂-depleted FG, and it has been computed as follows:(1) $ω_{FG−WW}^{{NH}_{3}} = ω_{FG−WW} (\frac{{\dot{m}}_{{CO}_{2}}^{storage}}{{\dot{m}}_{{NH}_{3}}^{FG−WW,in} - {\dot{m}}_{{NH}_{3}}^{FG−WW,out}})$ where $ω_{FG−WW}$ [ $MJ {kg}_{{CO}_{2} captured}^{- 1}$ ] is the equivalent work per unit of mass of CO₂ captured associated with the energy consumers of the FG-WW column and of the NH₃ desorber, ${\dot{m}}_{{CO}_{2}}^{storage}$ [ ${kg}_{{CO}_{2}} s^{−1}$ ] is the mass flowrate of CO₂ captured and sent to storage, and ${\dot{m}}_{{NH}_{3}}^{FG−WW,in}$ and ${\dot{m}}_{{NH}_{3}}^{FG−WW,out}$ [ ${kg}_{{NH}_{3}} s^{−1}$ ] are the mass flowrates of NH₃ in the FG stream entering and leaving, respectively, the FG-WW column.

Furthermore, the type of energy required is analysed, when relevant, by means of the breakdown of steam required in the reboilers of the CO₂ desorber, $q_{reb, {CO}_{2}}$ , the appendix stripper, $q_{reb,app}$ , and the NH₃ desorber, $q_{reb, {NH}_{3}}$ [ ${MJ}_{th} {kg}_{{CO}_{2} captured}^{- 1}$ ], vs. the electricity requirements—assumed to be equal to the specific equivalent work—for chilling, cooling, pumping and CO₂ compression, $e$ [ ${MJ}_{el} {kg}_{{CO}_{2} captured}^{- 1}$ ].

Besides the two types of energy requirements, the Specific Primary Energy Consumption per CO₂ Avoided (SPECCA) index, $ζ$ [ ${MJ}_{LHV} {kg}_{{CO}_{2} avoided}^{- 1}$ ], is used as energy performance indicator to assess advanced configurations aiming at improving the energetic performance of NH₃-based capture processes. As pointed out in our previous research [14], $ζ$ is preferred as energy performance indicator when the capture process is applied to a specific CO₂ point source, which might not have access to the electricity and steam produced in a power plant and where low temperature excess heat might be available for generation of (part of) the steam required in the capture plant. If the capture plant is an end-of-pipe technology that avoids modifications of the manufacturing process carried out in the CO₂ point source, as done in this work, $ζ$ can be computed as follows:(2) $ζ = \frac{q_{CCS}}{ϵ_{ref} - ϵ_{CCS}}$ where $q_{CCS}$ is the primary energy consumption of the Carbon Capture and Storage (CCS) plant, considering the Low Heating Value (LHV) of the fuel consumed, [ ${MJ}_{LHV}$ ], and $ϵ_{ref}$ and $ϵ_{CCS}$ are the CO₂ emissions of the point source without and with CCS plant, respectively, [ ${kg}_{{CO}_{2} avoided}$ ].

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The productivity depends on the volume of the pieces of equipment, which can be used as an indication of the trends of capital costs of the capture process. The cost of the CO₂ absorber has been reported as a major contributor to the investment costs associated with a solvent-based capture plant [55]. However, in this work we have added the productivity of the FG-WW column, $\Pr_{FG−WW}$ , to that of the CO₂ absorber, $\Pr_{{CO}_{2} abs}$ , in both cases referenced to the flowrate of CO₂ captured, i.e. [ ${kg}_{{CO}_{2} captured} m^{- 3} h^{- 1}$ ], in order to assess and compare advanced NH₃ abatement configurations. Therefore, the overall process productivity, $\Pr$ , is calculated as follows:(3) $\Pr = \frac{1}{\frac{1}{\Pr_{{CO}_{2} a b s}} + \frac{1}{\Pr_{FG−WW}}} = \frac{{\dot{m}}_{{CO}_{2}}^{storage}}{V_{{CO}_{2} abs} + V_{FG−WW}}$ where $V_{{CO}_{2} abs}$ and $V_{FG−WW}$ are the volumes of packing required in the CO₂ absorber and in the NH₃ absorber, respectively, which are calculated considering: (i) cylindrical shape, (ii) a hydraulic load at the bottom of each column at 70% of flooding [56], and (iii) a random packing of 25-mm Pall rings with a specific surface area of 207 m² m⁻³ used in the pilot tests reported by Yu et al. [52] and further detailed by Qi et al. [57] whose results were used to validate the rate-based model [14] used in this work. Additionally, another productivity related to the removal of NH₃ in the FG-WW column, $\Pr_{FG−WW}^{{NH}_{3}}$ [ ${kg}_{{NH}_{3} captured} m^{- 3} h^{- 1}$ ], has been defined:(4) $\Pr_{FG−WW}^{{NH}_{3}} = \frac{{\dot{m}}_{{NH}_{3}}^{FG−WW,in} - {\dot{m}}_{{NH}_{3}}^{FG−WW,out}}{V_{FG−WW}}$

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The consumption of water and the consumption of chemicals in aqueous solution, $F_{H_{2} O}^{in}$ and $F_{chem, i}^{in}$ , respectively, [ $kg t_{{CO}_{2} captured}^{- 1}$ ], with $i$ referring to NH₃ or to H₂SO₄, do not only contribute to the operating costs but can also be considered as an indication of process sustainability—in the case of aqueous NH₃ solution (25 wt%) and water—or as an indicator for the capture process retrofitability if chemicals that require stringent safety considerations are required for the capture process—in the case of aqueous H₂SO₄ solution (98 wt%).

Furthermore, the comparison among different advanced configurations has been performed at different ${\hat{c}}_{{NH}_{3}}$ values. Namely, 5, 7 and 9 ${mol}_{{NH}_{3}} {kg}_{H_{2} O}^{- 1}$ have been selected as representative of a low, a mid and a high typical concentration of solvent in NH₃-based capture processes [3], [14], [52]. The values of the remaining operating conditions for each ${\hat{c}}_{{NH}_{3}}$ value have been extracted from the optimization results obtained in our previous work [14] on the benchmark state-of-the-art CAP configuration shown in Fig. 1. The operating conditions have been selected aiming at minimizing the specific equivalent work for CO₂ capture efficiency equal to 0.9. These optimal values are indicated in Table 3, along with their corresponding process performance indicators.

The minimal specific equivalent work was obtained for the mid ${\hat{c}}_{{NH}_{3}}$ case, although the low ${\hat{c}}_{{NH}_{3}}$ case and the high ${\hat{c}}_{{NH}_{3}}$ case also performed very competitively in terms of energy consumption. On the one hand, increasing apparent NH₃ concentration in the solvent increases the NH₃ slip both in the CO₂-depleted FG leaving the CO₂ absorber, $y_{{NH}_{3}}^{FG−WW,in}$ , and in the CO₂ gas stream exiting the CO₂ desorber, $y_{{NH}_{3}}^{{CO}_{2} - WW,in}$ . As a result, there is an increase of the specific reboiler duties of the NH₃ desorber, $q_{reb, {NH}_{3}}$ , and of the appendix stripper, $q_{reb,app}$ , required for the recuperation of NH₃ from the vapour streams in the solvent recovery sections. Nevertheless, the main energy consumer of the process is the reboiler of the CO₂ desorber, $q_{reb, {CO}_{2}}$ , regardless of the NH₃ content of the solvent. On the other hand, operating the capture process at higher ${\hat{c}}_{{NH}_{3}}$ requires greater water make-up flowrates, $F_{H_{2} O}^{in}$ , to recover the NH₃ slipped to the gas streams, while the flowrates of aqueous NH₃ solution make-up, $F_{chem, {NH}_{3}}^{in}$ , and of aqueous H₂SO₄ solution make-up, $F_{chem, H_{2} {SO}_{4}}^{in}$ , remain almost unchanged. As far as the productivity of the process is concerned, increasing ${\hat{c}}_{{NH}_{3}}$ allows for greater CO₂ absorber productivity, $\Pr_{{CO}_{2} abs}$ , as pointed out earlier [14], while it has the opposite effect on the productivity of the FG-WW column, $\Pr_{FG−WW}$ . As a result, the overall process productivity, $\Pr$ , finds a maximum for the mid ${\hat{c}}_{{NH}_{3}}$ case, although the low ${\hat{c}}_{{NH}_{3}}$ case and the high ${\hat{c}}_{{NH}_{3}}$ also allowed for very competitive results. Notwithstanding that the selection of the optimal set of operating conditions that optimizes the capture process can be only determined by means of cost calculations, which is out of the scope of this work, the different effect of the three sets of operating conditions provided in Table 3 on the energy demand, on the absorbers’ size and on the NH₃ slip of each process section seems to make them ideal to carry out a general performance assessment and fair comparison of the advanced configurations proposed in this work.

Table 3. Process operating conditions that minimizes the specific equivalent work of the benchmark state-of-the-art CAP configuration shown in Fig. 1, aiming at 0.9 ${CO}_{2}$ capture efficiency, when fixing the apparent ${NH}_{3}$ concentration in the ${CO}_{2}$ -lean stream to 5 (low ${\hat{c}}_{{NH}_{3}}$ ), to 7 (mid ${\hat{c}}_{{NH}_{3}}$ ) and to 9 (high ${\hat{c}}_{{NH}_{3}}$ ) ${mol}_{{NH}_{3}} {kg}_{H_{2} O}^{- 1}$ , reported in the Supplementary Material of our previous optimization work [14]. The ${NH}_{3}$ slip in the ${CO}_{2}$ -depleted FG stream leaving the ${CO}_{2}$ absorber, $y_{{NH}_{3}}^{FG−WW,in}$ , and in the ${CO}_{2}$ gas stream exiting the ${CO}_{2}$ desorber, $y_{{NH}_{3}}^{{CO}_{2} - WW,in}$ , are also reported for each set of operating conditions, along with the corresponding values of the process performance indicators.

Empty Cell	Variable	Units	Simulation case
Empty Cell	Variable	Units	low ${\hat{c}}_{{NH}_{3}}$	mid ${\hat{c}}_{{NH}_{3}}$	high ${\hat{c}}_{{NH}_{3}}$
${CO}_{2}$ absorber parameters	${\hat{c}}_{{NH}_{3}}$	[ ${mol}_{{NH}_{3}} {kg}_{H_{2} O}^{- 1}$ ]	5.0	7.0	9.0
	$l_{lean}$	[ ${mol}_{{CO}_{2}} {mol}_{{NH}_{3}}^{- 1}$ ]	0.33	0.35	0.38
	$L^{lean} ∕ G^{in}$	[ $kg {kg}^{- 1}$ ]	5.5	5.0	5.5
	$f_{s}$	[-]	0.21	0.25	0.25
	$T_{pa}$	[ $° C$ ]	11	12	12
${CO}_{2}$ desorber parameters	$P_{{CO}_{2} des}$	[bar]	17.5	22.5	22.5
${CO}_{2}$ desorber parameters	$f_{cr}$	[-]	0.0500	0.0475	0.0450
FG-WW column parameters	${\hat{c}}_{{NH}_{3}}^{FG−WW}$	[ ${mol}_{{NH}_{3}} {kg}_{H_{2} O}^{- 1}$ ]	0.05	0.05	0.05
	${(L ∕ G)}^{FG−WW}$	[ $kg {kg}^{- 1}$ ]	0.1	0.1	0.1
	$T_{lean}^{FG−WW}$	[ $° C$ ]	1.5	1.5	1.5
	$T_{pa}^{FG−WW}$	[ $° C$ ]	1.5	1.5	1.5
	$f_{s}^{FG−WW}$	[-]	0.001	0.001	0.001
Internal ${NH}_{3}$ slip	$y_{{NH}_{3}}^{FG−WW,in}$	[ ${ppm}_{v}$ ]	3110	4780	7880
Internal ${NH}_{3}$ slip	$y_{{NH}_{3}}^{{CO}_{2} - WW,in}$	[ ${ppm}_{v}$ ]	610	1410	2050
Performance indicators	$ω$	[ $MJ {kg}_{{CO}_{2} captured}^{- 1}$ ]	1.057	1.043	1.078
	$ω_{FG−WW}$	[ $MJ {kg}_{{CO}_{2} captured}^{- 1}$ ]	0.031	0.037	0.045
	$ω_{FG−WW}^{{NH}_{3}}$	[ $MJ {kg}_{{NH}_{3} captured}^{- 1}$ ]	5.932	4.414	3.218
	$q_{reb, {CO}_{2}}$	[ ${MJ}_{th} {kg}_{{CO}_{2} captured}^{- 1}$ ]	2.22	2.16	2.23
	$T_{reb, {CO}_{2}}$	[ $° C$ ]	145	147	143
	$q_{reb,app}$	[ ${MJ}_{th} {kg}_{{CO}_{2} captured}^{- 1}$ ]	0.06	0.08	0.10
	$T_{reb,app}$	[ $° C$ ]	105	105	105
	$q_{reb, {NH}_{3}}$	[ ${MJ}_{th} {kg}_{{CO}_{2} captured}^{- 1}$ ]	0.11	0.14	0.17
	$T_{reb, {NH}_{3}}$	[ $° C$ ]	99	99	99
	$e$	[ ${MJ}_{el} {kg}_{{CO}_{2} captured}^{- 1}$ ]	0.287	0.277	0.290
	$\Pr$	[ ${kg}_{{CO}_{2} captured} m^{- 3} h^{- 1}$ ]	56.4	57.7	56.6
	$\Pr_{{CO}_{2} abs}$	[ ${kg}_{{CO}_{2} captured} m^{- 3} h^{- 1}$ ]	72.9	85.4	113.7
	$\Pr_{FG−WW}$	[ ${kg}_{{CO}_{2} captured} m^{- 3} h^{- 1}$ ]	249.2	178.0	112.6
	$\Pr_{FG−WW}^{{NH}_{3}}$	[ ${kg}_{{NH}_{3} captured} m^{- 3} h^{- 1}$ ]	1.314	1.476	1.574
	$F_{H_{2} O}^{in}$	[ $kg t_{{CO}_{2} captured}^{- 1}$ ]	7.04	14.26	27.52
	$F_{chem, {NH}_{3}}^{in}$	[ $kg t_{{CO}_{2} captured}^{- 1}$ ]	1.47	1.48	1.49
	$F_{chem, H_{2} {SO}_{4}}^{in}$	[ $kg t_{{CO}_{2} captured}^{- 1}$ ]	1.06	1.06	1.07

Unless mentioned otherwise, other assumptions regarding equipment specifications and utilities, required for process simulation and for the computation of the performance indicators, can be found in our previous work [14].

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